Hydrocarbon conversion processes using non-zeolitic molecular sieve catalysts

ABSTRACT

Hydrocarbon or oxygenate conversion process in which a feedstock is contacted with a non zeolitic molecular sieve which has been treated to remove most, if not all, of the halogen contained in the catalyst. The halogen may be removed by one of several methods. One method includes heating the catalyst in a low moisture environment, followed by contacting the heated catalyst with air and/or steam. Another method includes steam-treating the catalyst at a temperature from 400° C. to 1000° C. The hydrocarbon or oxygenate conversion processes include the conversion of oxygenates to olefins, the conversion of oxygenates and ammonia to alkylamines, the conversion of oxygenates and aromatic compounds to alkylated aromatic compounds, cracking and dewaxing.

This application is a continuation-in-part of U.S. application Ser. No.09/891,674, filed Jun. 25, 2001 and now U.S. Pat. No. 6,440,894.

FIELD OF THE INVENTION

The present invention relates to methods of removing halogen fromnon-zeolitic, molecular sieve catalysts, the catalysts produced fromsuch methods, and the use of such catalysts in hydrocarbon conversionprocesses.

BACKGROUND OF THE INVENTION

Molecular sieve catalysts used in a fluidized-bed reactor or a riserreactor will typically have an average particle diameter from 40 μm to300 μm. Catalyst particle size within this range is needed for properfluidization as well as to efficiently separate the catalyst from thegaseous products in a cyclone separator. To maintain the desiredcatalyst diameter the molecular sieve is formulated with othermaterials. Dilution of the molecular sieve with these materials is alsoused to control the rate of reaction, control the temperature of thereactor and regenerator, and to stabilize and protect the molecularsieve.

Formulated molecular sieve catalysts present a problem not found inother types of industrial catalysts, that is, how to maintain thephysical integrity of the molecular sieve catalyst during the fluidizedcyclic process of reaction, separation, and regeneration. The cycles ofreaction, separation, and regeneration are carried out at hightemperatures and high flow rates. Collisions and abrasions betweencatalyst particles, between the catalyst particles and reactor walls andbetween the catalyst particles and other parts of the unit tend to causephysical breakdown of the original catalyst into smaller catalystparticles known as fines. This physical breakdown is referred to ascatalyst attrition. The fines usually have particle diameters smallerthan 20 microns—much smaller than the original catalyst particles.Catalysts with higher attrition resistance are desirable because, amongother reasons, fewer fines are generated for disposal, lessenvironmental impact is caused by unrecoverable airborne particulates,optimal fluidized conditions are maintained, operating costs are lower,and less replacement catalyst is required.

Molecular sieve catalysts are formed by various methods, for example, byspray drying or extruding a slurry containing the molecular sieve andthe other catalyst components. The catalysts are formed by mixing thezeolitic molecular sieve with one or more binding agents such as one ormore types of alumina and/or silica. Matrix materials, typically clays,are also added and serve as diluents to control the rate of thecatalytic reaction, and to facilitate heat transfer during many stagesof the process. In U.S. Pat. No. 5,346,875 to Wachter et al. zeolite-Y(21.8 wt %) is mixed with Kaolin clay (14.5 wt %), silica sol (48.3 wt%), and Reheis chlorhydrol (15.4 wt %) to form a slurry which is thenspray dried and calcined. A conventional calcination procedure was used;heating at 550° C. in air for 2 hours.

Non-zeolitic, molecular sieve catalysts are known to convert oxygenates,particularly methanol, to light olefins. The oxygenate to olefin processincludes separate processing zones for conducting the catalyticreaction, product-catalyst separation, and catalyst regeneration. Theproduced olefin and other hydrocarbon products are separated from thecatalyst particles in a separator, suitably a cyclone separator. Aportion of the catalyst is recovered from the separator and passed to aregenerator. In the regenerator the non-zeolitic molecular sievecatalyst contacts a combusting gas, e.g. air, at a temperaturesufficient to burn off carbon deposits, commonly referred to a coke,that accumulate on the surface and in the pores of the catalyst. Theregenerated catalyst is then returned to the oxygenate conversionreactor.

In this process, the non-zeolitic molecular sieve catalyst is subjectedto great mechanical stresses. As the catalyst is transferred from thereaction zone to cyclone separators, to regenerators, and finally backto the reaction zone the catalyst will tend to disintegrate intocatalyst fines. These catalyst fines must be removed from the reactorprocess and discarded. No matter how resistant the catalyst is toattrition, eventually the oxygenate to olefin process will break downthe non-zeolitic molecular sieve catalyst because the catalyst movesthrough the system at such high speeds. The resistance of the catalystto attrition is an important property of the catalyst.

In PCT Publication No. WO 99/21651 to Wachter et al. and U.S. Pat. No.4,973,792 to Lewis et al., silicoaluminophosphate (SAPO) molecular sievecatalysts were produced by preparing a slurry containing SAPO-34, Kaolinclay, and Reheis chlorhydrol. The slurry was then directed to a spraydryer to form catalyst particles with the desired size. The spray driedcatalysts were calcined, however the conditions of the calcination werestated to be not critical.

In U.S. Pat. Nos. 5,248,647 and 5,095,163 to Barger et al. SAPOmolecular sieve is mixed with an aqueous silica sol and spray dried. Thespray dried catalyst is mixed with an aqueous solution of ammoniumsulfate at 60° C. three times, then washed with water and dried at 100°C. The dried, ion-exchanged catalyst is then calcined in air at 550° C.for over 3.3 hours and then the temperature is lowered to ambient roomtemperature over a period of 2 hours. A portion of this catalyst is thencontacted with steam at 725° C. or 750° C. for 10 hours. Steam treatmentfollowing calcination is shown to increase catalyst life, increaseselectivity to ethylene and propylene, and decrease selectivity topropane.

If SAPO molecular sieve catalysts are ever going to be used commerciallyto convert oxygenates to olefins, catalysts with greater attritionproperties are needed. For this reason, the Applicants' sought todevelop SAPO catalysts with a relatively high resistance to attrition.

SUMMARY OF THE INVENTION

The present invention is directed to methods of removing a portion ofthe halogen present in non-zeolitic molecular sieve catalysts. Oneembodiment of removing halogen includes heating the catalyst in a lowmoisture environment at a temperature from about 400° C. to about 1000°,and contacting the heated catalyst with steam at a temperature fromabout 400° C. to about 1000° C. to produce a steam-treated catalyst.Preferably, the low moisture environment contains less than 5% byvolume, more preferably less than 1% by volume, water. The steamtreatment can take place in an oxygen environment. Also, it is preferredthat the steam treatment take place in an environment containing atleast 10% by volume water. In the preferred embodiment, the steamtreatment can remove from about 50% to about 99% by weight, morepreferably from about 90% to about 99% by weight, of halogen from theheated catalyst. The method can be used to remove halogen fromsilicoaluminophosphate and/or aluminophosphate molecular sieve selectedfrom the group consisting of SAPO-5, SAPO-8, SAPO-11, SAPO-16, SAPO-17,

SAPO-18, SAPO-20, SAPO-31, SAPO-34, SAPO-35, SAPO-36, SAPO-37, SAPO-40,SAPO-41, SAPO-42, SAPO-44, SAPO-47, SAPO-56, ALPO-5, ALPO-11, ALPO-18,ALPO-31, ALPO-34, ALPO-36, ALPO-37, ALPO-46, the metal containing formsof each thereof, or mixtures thereof.

In another embodiment, a portion of the halogen can be removed from anon zeolitic molecular sieve catalyst by heating the catalyst in anoxygen environment at a temperature from about 400° C. to about 1000° C.to produce a heated catalyst, and contacting the heated catalyst withsteam at a temperature from about 400° C. to about 1000° C. Preferably,the oxygen environment contains greater than about 10% by volume oxygen.It is also preferred, that the steam treatment take place in anenvironment containing at least about 10% by volume water. In manycases, the halogen to be removed will be chlorine, and preferably fromabout 70% to about 99% by weight, more preferably from about 90% toabout 99% by weight, of the chlorine will be removed from the heatedcatalyst.

In another embodiment, a portion of the halogen can be removed from anon zeolitic molecular sieve catalyst by calcining the catalyst in anenvironment containing steam at a temperature from about 400° C. toabout 1000° C., preferably from about 500° C. to about 800° C., and morepreferably from about 550° C. to about 700° C., to remove from about 70%to about 99.99% by weight of the halogen from the catalyst. If thehalogen to be removed from the catalyst is chlorine, the likely sourceof the chlorine is aluminum chlorhydrol that is used to produce thecatalyst. The environment can contain from 5% to about 10% by volumewater, or at least 10% by volume, water. The environment can furthercontain air, nitrogen, helium, flue gas, or any combination thereof.

In one embodiment, the catalyst is heated in a low-moisture environmentat a temperature of from about 400° C. to about 1000° C. to remove atleast about 50% by weight of the halogen in the catalyst prior to steamtreatment. Preferably, the low moisture environment contains less thanabout 5% by volume, more preferably less then about 1% by volume, water.Also, the steam-treated catalyst can optionally be heated in an oxygenenvironment that contains greater than about 5% by volume oxygen.

In another embodiment, a portion of the halogen can be removed from asilicoaluminophosphate molecular sieve catalyst by heating the catalystin a low moisture environment at a temperature from 400° C. to about1000° C. to remove at least about 50% by weight of the chlorine from thecatalyst, followed by contacting the heated catalyst in a secondcalcination environment containing about 5% to about 10% by volume waterat a temperature from 400° C. to about 1000° C. Preferably, the lowmoisture environment contains less than about 1% by volume water.

The invention is also directed to a catalyst containing a non zeoliticmolecular sieve, inorganic oxide matrix, and matrix material, whereinthe catalyst contains from about 10 ppmw to about 600 ppmw by weighthalogen. Generally, the halogen is chlorine, and the catalyst willcontain from about 10 ppmw to about 200 ppmw, preferably from about 10ppmw to about 80 ppmw, chlorine. It is also preferred that the catalysthave a GAL Index of less than about 5, more preferably less than about3, most preferably less than about 2. The non-zeolitic molecular sievein the catalyst is preferably selected from SAPO-5, SAPO-8, SAPO-11,SAPO-16, SAPO-17, SAPO-18, SAPO-20, SAPO-31, SAPO-34, SAPO-35, SAPO-36,SAPO-37, SAPO-40, SAPO-41, SAPO-42, SAPO-44, SAPO-47, SAPO-56, ALPO-5,ALPO-11, ALPO-18, ALPO-31, ALPO-34, ALPO-36, ALPO-37, ALPO-46, the metalcontaining forms of each thereof, or mixtures thereof. Preferably, thecatalyst contains about 20% to about 45% by weight, more preferably fromabout 25% to about 42% by weight, non-zeolitic molecular sieve, about 5%to about 20% by weight, more preferably about 8% to about 15% by weight,of inorganic oxide matrix, and about 20% to about 70% by weight, morepreferably from about 40% to about 60% by weight, matrix material. Inthe preferred embodiment, the inorganic oxide matrix contains analuminum oxide matrix that is formed from the heat treatment of aluminumchlorhydrol.

The present invention also relates to hydrocarbon and oxygenateconversion processes in which a feedstock is contacted with a nonzeolitic catalyst from which halogen has been removed. Morespecifically, the present invention relates to a hydrocarbon oroxygenate conversion process comprising the steps of (a) introducing afeedstock to a reactor system in the presence of a catalyst comprising anon zeolitic molecular sieve, inorganic oxide matrix, and matrixmaterial, wherein the catalyst contains from 10 ppm to 600 ppm by weighthalogen; (b) withdrawing from the reactor system an effluent stream; and(c) passing the effluent gas through a recovery system recovering atleast the one or more conversion products.

In another embodiment, the present invention relates to a hydrocarbon oroxygenate conversion process comprising the steps of: (a) providing ahalogen-containing non-zeolitic molecular sieve catalyst; (b) heatingdie catalyst in a low moisture environment at a temperature from 400° C.to 1000° C.; (c) contacting the heated catalyst with steam at atemperature from 400° C. to 1000° C. to produce a steam-treatedcatalyst; (d) introducing a feedstock to a reactor system in thepresence of the steam-treated catalyst obtained at step (c); (e)withdrawing from the reactor system an effluent stream; and (f) passingthe effluent gas through a recovery system recovering at least the oneor more conversion products.

In another embodiment, the present invention relates to a hydrocarbon oroxygenate conversion process comprising the steps of (a) providing ahalogen-containing non-zeolitic molecular sieve catalyst; (b) heatingthe catalyst in an oxygen environment at a temperature from 400° C. to1000° C.; (c) contacting the heated catalyst with steam at a temperaturefrom 400° C. to 1000° C. to produce a steam-treated catalyst; Cd)introducing a feedstock to a reactor system in the presence of thesteam-treated catalyst obtained at step (c); (e) withdrawing from thereactor system an effluent stream; and (f) passing the effluent gasthrough a recovery system recovering at least the one or more conversionproducts.

In a further embodiment, the present invention relates to a hydrocarbonor oxygenate conversion process comprising the steps of (a) providing ahalogen-containing non-zeolitic molecular sieve catalyst; (b) heatingthe catalyst in an environment containing steam at a temperature from400° C. to 1000° C.; (c) removing from 70% to 99.99% by weight of thehalogen from the catalyst, thereby producing a steam-treated catalyst;(d) introducing a feedstock to a reactor system in the presence of thesteam-treated catalyst obtained at step (c); (e) withdrawing from thereactor system an effluent stream; and (f) passing the effluent gasthrough a recovery system recovering at least the one or more conversionproducts.

In each of the preceding embodiments, the present invention isapplicable to a wide range of processes including those in which thefeedstock comprises one or more oxygenates, ammonia, aromatic compounds,or mixtures thereof, which are converted to olefins, alkylamines oralkylated aromatic compounds. The invention is also applicable forfeedstock cracking and dewaxing.

BRIEF DESCRIPTION OF THE DRAWING

The present invention will be better understood by reference to theDetailed Description of the Invention when taken together with theattached drawing, wherein FIG. 1 is a schematic representation of oneembodiment for removing chlorine from a formed catalyst.

DETAILED DESCRIPTION OF THE INVENTION

To produce non-zeolitic molecular sieve catalyst with a relatively highresistance to attrition, an inorganic oxide sol that contains halogencan be used. A preferred route to produce non-zeolitic molecular sievecatalyst is to use an alumina sol that contains chlorine, morepreferably aluminum chlorhydrol, as a binder. The inorganic oxide solfunctions as a “glue” which binds the catalyst components together.However, using an inorganic oxide sol that contains halogen presents aproblem not associated with the use of halogen-free binders. A portionof the halogen from the inorganic oxide sol remains in the formedcatalyst. It is desirable to remove most, if not nearly all, of thehalogen from the catalyst before the catalyst is used in the oxygenateto olefin process. If most of the halogen is not removed from thecatalyst, halogen-containing acids will form in the oxygenate to olefinreactor. Over time, the released acid will corrode the oxygenate toolefin reactor and other process units. While the invention will befurther illustrated for the case where the halogen is chlorine, itshould be understood that the invention applies to other halogens aswell, such as fluorine, bromine and iodine. In the case of a catalystcontaining chlorine, hydrochloric acid will form in the oxygenate toolefin reactor. HCl may be in the gas or condensed form, usually in ahydrated form, hereinafter referred to as HCl_((aq)). All forms of acidsare potentially corrosive, the hydrated form being the most corrosive.

The invention addresses the problem associated with the use of inorganicoxide sols that contains halogen by removing much of the halogen fromthe catalyst during calcination of the catalyst. The invention addressesthese problems by providing methods of heat treating or calcining aformed non-zeolitic molecular sieve catalyst prepared with an inorganicoxide sol that contains halogen. The methods of the invention minimizethe production of halogen-containing acids, or at least confines much ofthe produced halogen-containing acids to a single heating or calcinationunit that can be designed to accommodate the corrosive effects ofhalogen-containing acids. The methods of the invention also reduce theamount of halogen remaining in the catalyst over that of conventionalprocedures.

The catalyst is made by preparing a slurry containing non-zeoliticmolecular sieve, an inorganic oxide binder, and a matrix material. Theslurry is then dried and shaped in a forming unit. Preferably, theslurry is spray dried, and a dry powder catalyst with an averagecatalyst particle size is obtained. The formed catalyst is then heattreated, i.e., calcined.

Calcination is used to remove the template molecule from the cagestructure of the framework. During calcination all or part of thetemplate molecule exits the cage structure. Calcination is also used toharden the formed catalyst particle. The relatively high temperaturesused during calcination transform the inorganic oxide sol to aninorganic oxide matrix. It is this inorganic oxide matrix that increasesthe attrition resistance of the catalyst particle.

If a conventional calcination procedure is used on a catalyst containingchlorine, that is, heating in air at temperatures greater than 400° C.,large amounts of HCl are produced over time in the calcination unit. Theformation of HCl_((aq)) is the result of small amounts of water or watervapor contained in the air and the water generated from the oxidativecombustion of the organic template during calcination. The released HCl,if not accounted for, will eventually corrode the heating or calcinationunit. Therefore, it is desirable to control the removal of chlorine fromthe catalyst in a manner that will either minimize the amount of HClproduced during the calcination process or limit the evolution of HCl toa single calcination unit.

A conventional calcination procedure also does not remove enough of thehalogen from the catalyst. In the case of chlorine, the remainingchlorine in the catalyst is then released into the oxygenate to olefinreactor and other oxygenate to olefin process units as HCl_((aq)) due tothe hydrothermal conditions of the oxygenate to olefin process. If notaccounted for, the release of this HCl_((aq)) will corrode the oxygenateto olefin process units. The presence of HCl_((aq)) in the olefinmonomer feed used for polymerization might also damage or poisonexpensive polymerization catalysts. Therefore, it is desirable to removeas much chlorine from the catalyst during the calcination process so asto minimize the amount of HCl_((aq)) released into the oxygenate toolefin process units.

As a result of using the calcination methods of the invention, anon-zeolitic molecular sieve catalyst with low amounts of halogen isobtained. A preferred catalyst of the invention contains a SAPOmolecular sieve, an aluminum oxide matrix, and clay, most preferablyKaolin. The catalyst will also contain some halogen resulting from theuse of a binder that contains halogen. Although the invention isdirected to removing as much halogen from the catalyst as efficientlypossible, some of the halogen is not removed during the calcinationprocess. Following the calcination procedures of the invention, thecatalyst will contain from about 10 ppmw to 600 ppmw halogen, preferablyfrom about 10 ppmw to 200 ppmw halogen, more preferably from about 10ppmw to 60 ppmw halogen. The catalyst will also have a Gross AttritionLoss (GAL) Index of less than 5, preferably a GAL Index less than 3,more preferably a GAL Index less than 2. The smaller the GAL Index, themore resistant to attrition is the catalyst.

Non Zeolitic Molecular Sieve

The catalyst used according to the present invention contains a nonzeolitic molecular sieve. Examples of suitable non-zeolitic molecularsieves are silicoaluminophosphates (SAPOs) and aluminophosphates(ALPOs). In general, SAPO molecular sieves comprise a molecularframework of corner-sharing [SiO₄], [AlO₄], and [PO₄] tetrahedral units.The [PO₄] tetrahedral units are provided by a variety of compositions.Examples of these phosphorus-containing compositions include phosphoricacid, organic phosphates such as triethyl phosphate, andaluminophosphates. The [AlO₄] tetrahedral units are provided by avariety of compositions. Examples of these aluminum-containingcompositions include aluminum alkoxides such as aluminum isopropoxide,aluminum phosphates, aluminum hydroxide, sodium aluminate, andpseudoboehmite. The [SiO₄] tetrahedral units are provided by a varietyof compositions. Examples of these silicon-containing compositionsinclude silica sols and silicium alkoxides such as tetra ethylorthosilicate. The phosphorus-, aluminum-, and silicon-containingcompositions are mixed with water and a template molecule and heatedunder appropriate conditions to form the molecular sieve.

SAPO molecular sieves are generally classified as being microporousmaterials having 8, 10, or 12 membered ring structures. These ringstructures can have an average pore size ranging from about 3.5–15angstroms. Preferred are the small pore SAPO molecular sieves having anaverage pore size of less than about 5 angstroms, preferably an averagepore size ranging from about 3.5 to 5 angstroms, more preferably from3.5 to 4.2 angstroms. These pore sizes are typical of molecular sieveshaving 8 membered rings.

An aluminophosphate (ALPO) molecular sieve can also be included in thecatalyst composition. Aluminophosphate molecular sieves are crystallinemicroporous oxides which can have an AlPO₄ framework. They can haveadditional elements within the framework, typically have uniform poredimensions ranging from about 3 Angstroms to about 10 Angstroms, and arecapable of molecular size selective separations of molecular species.More than two dozen structure types have been reported, includingzeolite topological analogues.

For a catalyst used in the conversion of oxygenate to light olefin thenon-zeolitic molecular sieve will have a relatively low Si/Al₂ ratio. Ingeneral, for SAPOs, a Si/Al₂ ratio of less than 0.65 is desirable, witha Si/Al₂ ratio of not greater than 0.40 being preferred, and a Si/Al₂ratio of not greater than 0.32 being particularly preferred. A Si/Al₂ratio of not greater than 0.20 is most preferred.

Substituted SAPOs and ALPOs can also be used in this invention. Thesecompounds are generally known as MeAPSOs, MeAPOs, metal-containingsilicoaluminophosphates or metal-containing aluminophosphates. The metalcan be alkali metal ions (Group IA), alkaline earth metal ions (GroupIIA), rare earth ions (Group IIIB, including the lanthanide elements,and the additional transition cations of Groups IB, IIB, IVB, VB, VIB,VIIB, and VIIIB. Preferably, the Me represents atoms such as Zn, Ni, andCu. These atoms can be inserted into the tetrahedral framework through a[MeO₂] tetrahedral unit. Incorporation of the metal component istypically accomplished by adding the metal component during synthesis ofthe molecular sieve. However, post-synthesis metal incorporation canalso be used.

SAPO and ALPO molecular sieves that can be used include SAPO-5, SAPO-8,SAPO-11, SAPO-16, SAPO-17, SAPO-18, SAPO-20, SAPO-31, SAPO-34, SAPO-35,SAPO-36, SAPO-37, SAPO-40, SAPO-41, SAPO-42, SAPO-44, SAPO-47, SAPO-56,ALPO-5, ALPO-11, ALPO-18, ALPO-31, ALPO-34, ALPO-36, ALPO-37, ALPO-46,the metal containing forms thereof, and mixtures thereof. Preferred areSAPO-18, SAPO-34, SAPO-35, SAPO-44, SAPO-56, ALPO-18 and ALPO-34,particularly SAPO-18, SAPO-34, ALPO-34 and ALPO-18, including the metalcontaining forms thereof, and mixtures thereof. As used herein, the termmixture is synonymous with combination and is considered a compositionof matter having two or more components in varying proportions,regardless of their physical state.

SAPO and ALPO molecular sieves are synthesized by hydrothermalcrystallization methods generally known in the art. See, for example,U.S. Pat. Nos. 4,440,871; 4,861,743; 5,096,684; and 5,126,308, thedisclosures of which are fully incorporated herein by reference. Areaction mixture is formed by mixing together reactive silicon, aluminumand phosphorus components, along with at least one template. Generallythe mixture is sealed and heated, preferably under autogenous pressure,to a temperature of at least 100° C., preferably from 100–250° C., untila crystalline product is formed.

Formation of the crystalline product can take anywhere from around 2hours to as much as 2 weeks. In some cases, stirring or seeding withcrystalline material will facilitate the formation of the product.Typically, the molecular sieve product is formed in solution. It can berecovered by standard means, such as by centrifugation or filtration.The product can also be washed, recovered by the standard means, anddried. In one method, the molecular sieve is washed and collected by afiltration process that maintains the molecular sieve in slurry form.This process includes adding wash fluid as the molecular sieve isconcentrated from the synthesis solution.

Additional molecular sieve materials can be included as a part of thenon zeolitic catalyst or they can be used as separate molecular sievecatalysts in admixture with the non zeolitic molecular sieve catalyst ifdesired. Structural types of small pore molecular sieves that aresuitable for use in this invention include AEI, AFT, APC, ATN, ATT, ATV,AWW, BIK, CAS, CHA, CHI, DAC, DDR, EDI, ERI, GOO, KFI, LEV, LOV, LTA,MON, PAU, PHI, RHO, ROG, THO, and substituted forms thereof. Structuraltypes of medium pore molecular sieves that are suitable for use in thisinvention include MFI, MEL, MTW, EUO, MTT, HEU, FER, AFO, AEL, TON, andsubstituted forms thereof. These small and medium pore molecular sievesare described in greater detail in the Atlas of Zeolite StructuralTypes, W. M. Meier and D. H. Olsen, Butterworth Heineman, 3rd ed., 1997,the detailed description of which is explicitly incorporated herein byreference. Preferred molecular sieves which can be combined with asilicoaluminophosphate and/or an aluminophosphate catalyst includeZSM-5, ZSM-34, erionite, and chabazite.

Binders

Once the desired type or types of non-zeolitic molecular sieve isselected based upon the desired activity and selectivity of thecatalyst, other materials are blended with the non-zeolitic molecularsieve. One of these materials includes one or more binders, such as atype of hydrated alumina, and/or an inorganic oxide sol such as aluminumchlorhydrol. The inorganic oxide sol is essentially a “glue” which bindsthe catalyst components together upon thermal treatment. After theformed catalyst particle is formed and heated, the inorganic oxide solis converted to an inorganic oxide matrix component. For example, analumina sol will convert to an aluminum oxide matrix following a heattreatment of the formed catalyst. Aluminum chlorhydrol is a hydroxylatedaluminum based sol containing chloride as the counter ion. Aluminumchlorhydrol has the general formula of Al_(m)O_(n)(OH)_(o)Cl_(p).xH₂Owherein m is 1 to 20, n is 1 to 8, o is 5 to 40, p is 2 to 15, and x is0 to 30. Although the equilibria of alumina species in the sol iscomplex, the predominant species is believed to be[Al₁₃O₄(OH)₂₄Cl₇(H₂O)₁₂]. In addition, other alumina materials may beadded with the aluminum chlorhydrol. Materials that can be used include,but are not necessarily limited to aluminum oxyhydroxide, γ-alumina,boehmite, diaspore, and transitional aluminas such as α-alumina,β-alumina, γ-alumina, δ-alumina, ε-alumina, κ-alumina, and ρ-alumina.Aluminum trihydroxide, such as gibbsite, bayerite, nordstrandite,doyelite, and mixtures thereof, also can be used. A sufficient amount ofthe binder is added to the slurry mixture so that the amount of theresultant inorganic oxide matrix in the catalyst, not including theinorganic oxide framework of the non-zeolitic molecular sieve, is fromabout 2% to about 30% by weight, preferably from about 5% to about 20%by weight, and more preferably from about 7% to about 12% by weight.

Matrix Materials

The non zeolitic molecular sieve catalysts will also contain clay,preferably Kaolin. Matrix materials may also include compositions suchas various forms of rare earth metals, metal oxides, titania, zirconia,magnesia, thoria, beryllia, quartz, silica or silica or silica sol, andmixtures thereof. The added matrix materials components are effective inreducing, inter alia, overall catalyst cost, acting as a thermal sink toassist in heat shielding the catalyst during regeneration, densifyingthe catalyst and increasing catalyst strength. The use of matrixmaterials such as naturally occurring clays, e.g., bentonite and kaolin,improves the crush strength of the catalyst under commercial operatingconditions. Thus, the addition of clays improve upon the attritionresistance of the catalyst. The inactive materials also serve asdiluents to control the rate of conversion in a given process so thatmore expensive means for controlling the rate of reaction is eliminatedor minimized. Naturally occurring clays which can be used in the presentinvention include the montmorillonite and kaolin families which includethe sabbentonites, and the kaolins, commonly known as Dixie, McNamee,Georgia and Florida clays, or other in which the main mineralconstituent is haloysite, kaolinite, dickite, nacrite, or anauxite.

As with most catalysts clay is used in the invention as an inertdensifier, and for the most part the clay has no effect on catalyticactivity or selectivity. Kaolin's ability to form pumpable, high solidcontent slurries, low fresh surface area, and ease of packing because ofits platelet structure makes it particularly suitable for catalystprocessing. The preferred average particle size of the kaolin is 0.1 μmto 0.6 μm with a D90 particle size of about 1 μm. Because ofenvironmental concerns, the crystalline silica content of the clay hasalso become an important parameter.

Mixing and Spray Drying.

Rigorous mixing of the catalyst components is necessary to produce ahard, dense, homogeneous catalyst particle. The primary consequence ofpoor mixing are poor attrition and poor catalyst density. Stratificationof the components caused by incomplete mixing can also effect theactivity and selectivity of the catalyst. Generally, the mixers are of ahigh shear type because of the thixotropic nature of the slurries. Theresultant slurry may be colloid-milled for a period sufficient to obtaina desired sub-particle texture, sub-particle size, and/or sub-particlesize distribution.

The catalyst particle contains a plurality of catalyst sub-particles.The average diameter of the catalyst particle is from 40 μm to 300 μm,preferably from 50 μm to 200 μm. The catalyst sub-particles containnon-zeolite molecular sieve, typically SAPO molecular sieve, an aluminumoxide matrix, and a matrix material, typically clay. Preparation of thecatalyst begins with mixing one or more non-zeolite molecular sieve, oneor more inorganic oxide sols, one or more matrix materials, and a fluid,typically water, to form a slurry. Other fluids, e.g., alcohol, can beused along with the water.

The preferred slurry is prepared by mixing the non-zeolitic molecularsieve with aluminum chlorhydrol and Kaolin clay, together or insequence, in dry form or as slurries. If the solids are added togetheras dry solids, a limited and controlled amount of water is added. Theslurry may also contain other materials including other forms ofmolecular sieve, other binders, and other matrix materials. Themesoporosity of the catalyst and the mechanical strength of the catalystis dependent on the amount of water contained in the slurry. In general,it has been found that the weight percent of solids in the slurry canrange from 20% to 70% by weight, preferably from 40% to 60% by weight.When the weight percent of solids in the slurry is greater than 70% byweight, the viscosity of the slurry is too high to spray dry, and whenthe weight percent of solids in the slurry is less than 20% by weightthe attrition resistance of the catalyst is poor. It is also desirablethat the density of the slurry be greater than 1.1 g/cc, and preferablygreater than 1.18 g/cc to form the catalysts of this invention.

The solid content of the slurry will contain about 10% to about 50%,preferably from about 20% to about 45% by weight, non-zeolitic molecularsieve, about 5% to about 20%, preferably from about 8% to about 15% byweight, binder, and about 30% to about 80%, preferably about 40% toabout 60% by weight, matrix material. The slurry is mixed or milled toachieve a sufficiently uniform slurry of catalyst sub-particles. Theslurry is then fed to a forming unit to produce catalyst particles. Theforming unit is maintained at a temperature sufficient to remove most ofthe water from the formed catalyst particles. Preferably, the formingunit is a spray dryer. The formed catalyst particles typically take theform of microspheres. Typically, the slurry is fed to a spray drier atan average inlet temperature ranging from 200° C. to 450° C., and anoutlet temperature ranging from 100° C. to about 225° C.

During spray drying, the slurry is passed through a nozzle whichdistributes the slurry into small droplets, resembling an aerosol. Asingle nozzle unit or multiple nozzle unit may be used to disperse aninlet stream of slurry (single-fluid nozzle) into the atomizationchamber. Alternatively, a multiple nozzles may be used to co-feed theslurry into the atomization chamber. Alternatively, the slurry isdirected to the perimeter of a spinning wheel which also distributes theslurry into small droplets. The size of the distributed slurry dropletsis controlled by many factors including flow rate, pressure, andtemperature of the slurry, the shape and dimension of the nozzle(s), orthe spinning rate of the wheel. The droplets are then dried in aco-current or counter-current flow of air passing through the spraydrier. Dry catalyst particles in the form of a powder are recovered fromeach droplet.

Catalyst particle size to some extent is controlled by the solidscontent of the slurry and its viscosity. All else being equal, thecatalyst particle size is directly proportional to the solids content ofthe slurry. However, control of the catalyst particle size and sphericalcharacteristics also depend on the size and shape of the drying chamberas well as the atomization procedure used. A Boltzmann distribution ofcatalyst particle size is invariably obtained around a mean, which isusually set at approximately 70 μm average catalyst particle size. Theaverage catalyst particle size is controlled by a variation in theslurry feed properties to the dryer and by the conditions ofatomization. It is preferred that the formulated catalyst compositionhave a catalyst size from 40 μm to 300 μm, more preferably 50 μm to 200μm, most preferably 50 μm to 150 μm.

Calcination.

To harden and/or activate the formed catalysts a heat treatment, i.e.,calcination, at an elevated temperature is usually necessary.Ordinarily, catalysts with alumina or silica binders are heated in acalcination environment at a temperature between 500° C. and 800° C. Theconventional calcination environment is air, which may include smallamounts of water vapor.

The invention provides methods of heat treating a formed non-zeoliticmolecular sieve catalyst prepared with an inorganic oxide sol thatcontains halogen. The methods of the invention minimize the productionof halogen-containing acids or at least confines much of the producedhalogen-containing acids to a single heating unit. The schematic diagramin FIG. 1 depicts one embodiment of the invention, in which achlorine-containing SAPO catalyst 12 is used by way of example. Catalyst12 is supplied from a forming unit, preferably a spray dryer, and isdirected to a heat treatment unit 14. The catalyst is heated at atemperature from about 400° C. to about 1000° C., preferably from about500° C. to about 800° C., most preferably from about 550° C. to about700° C. in a low moisture calcination environment containing less than5% by volume water, preferably less than 1% by volume water. The lowmoisture calcination environment can be provided by using a dry gas 18,e.g., air that has been adequately dried, nitrogen, helium, flue gas, orany combination thereof. In the preferred embodiment, the catalyst isheated in nitrogen off gas at a temperature from about 600° C. to about700° C. Nitrogen off gas is the gas produced from the boil-off gas of aliquid nitrogen source. Heating is carried out for a period of timesufficient to remove chlorides, typically for a period of from 0.5 to 10hours, preferably of from 1 to 5 hours, most preferably from 2 to 4hours.

As the catalyst is heated in the low moisture calcination environmentmost of the chlorine is removed as chlorine gas or as a non-hydratedform of hydrochloric acid (HCl_((g))). HCl_((g)) is not as corrosive asHCl_((aq)). Approximately 60% to 98% by weight, preferably 85% to 98% byweight, of the chlorine in the formed catalyst 12 is removed during theheat treatment in the low moisture calcination environment. Followingthis heat treatment the catalyst 16 contains less than about 6000 ppmwchlorine, preferably less than about 3000 ppmw chlorine.

Following the low moisture heat treatment, the catalyst 16 is directedto heating unit 20. The catalyst in heating unit 20 is heated in asecond calcination environment 24. This second calcination environment24 contains from about 5% to about 10% by volume water. The remainingvolume of gas in the calcination environment 24 may include air,nitrogen, helium, flue gas, or any combination thereof. The second heattreatment of the catalyst 16 will take place at a temperature from about400° C. to about 1000° C., preferably from about 500° C. to about 800°C., more preferably from about 600° C. to about 700° C. The periodduring which the catalysts is heated in unit 20 ranges from 0.1 to 5hours, preferably from 0.25 to 4 hours. This second heat treatmentresults in a loss of about 2% to about 95% of the chlorine remaining incatalyst 16. Catalyst 22 will contain less than about 600 ppmw chlorine,preferably less than about 200 ppmw chlorine, more preferably less thanabout 80 ppmw chlorine.

In another embodiment, the second calcination environment contains atleast 10% by volume water. The remaining volume of gas in the secondcalcination environment may include air, nitrogen, helium, flue gas, orany combination thereof. Preferably, the second calcination environmentcontains air. A catalyst that is contacted with a calcinationenvironment containing at least 10% by volume water is said to besteam-treated. Steam treatment results in a loss of about 50% to 99%,preferably in a loss of about 90% to about 99% of the chlorine remainingin the catalyst following the low moisture heat treatment. Thesteam-treated catalyst will contain about 10 ppmw to about 400 ppmwchlorine, preferably about 10 ppmw to about 200 ppmw chlorine, morepreferably about 10 ppmw to about 80 ppmw chlorine.

Steam treatment of the catalyst will take place at a temperature fromabout 400° C. to about 1000° C., preferably from about 500° C. to about800° C., more preferably from about 600° C. to about 700° C. The periodduring which the catalysts is heated in unit 20 ranges from 0.1 to 5hours, preferably from 0.25 to 4 hours. Although temperatures of about400° C. are sufficient to adequately remove most of the chlorine fromthe catalyst, the rate at which the additional chlorine is removed willbe lower than if a higher temperature, e.g., 600° C., is used duringsteam treatment. On the other hand, if the temperature of the steamtreatment is too high, e.g., greater than 1000° C., degradation of thecatalyst may occur. The temperature at which degradation of the catalystwill occur will vary for different catalyst formulations and variousnon-zeolitic molecular sieve.

The low moisture heat treatment followed by steam treatment can removeabout 70% to about 99.99% by weight, preferably about 95% to about99.99% by weight, more preferably about 98% to about 99.99% by weight,of the chlorine in the formed catalyst. The steam treatment will produceHCl_((aq)), but the amount of HCl_((aq)) produced is significantlyreduced because most of the chlorine is removed during the initial heattreatment in the low moisture calcination environment. As a result, theproduction of the HCl_((aq)) is minimized. Also, if separate heatingunits are used the production of HCl_((aq)) will be confined to thesteam treatment unit, which can be designed to accommodate theHCl_((aq)) produced.

If air is not used in the steam treatment, the catalyst may be calcinedin a calcination environment containing at least 3% by volume,preferably at least 10% by volume, oxygen to remove template materialthat may have remained in the pores of the sieve. A catalyst that hasbeen calcined in an environment that contains at least 3% by volumeoxygen is said to be oxygen treated. The oxygen environment may beprovided by air or a mixture of air and nitrogen. The calcinationtemperature of this oxygen environment may be the same or different thanthe temperature of the steam treatment.

It is to be understood that although FIG. 1 depicts more than oneheating unit for each type of heat treatment, a single heating unit maybe used. In this case, the heating environment is changed by alternatingthe type of gas flow, e.g., from nitrogen off gas to steam, or from airto steam. Alternatively, different heating zones in a singular heatingunit may be used according to the invention. Each heating zone willcontain a different calcination environment with a transition zonedisposed between the heating zones. The temperature and gas flow foreach heating zone or heating unit can be programmed to minimize the timerequired to remove the desired amount of chlorine, while minimizing theamount of HCl_((aq)) produced. The heat and steam treatments may be donein any of a number of heating units well known to those skilled in theart including moving bed reactors, rotary kilns, rotary calciners,fluidized beds and packed-bed batch reactors.

In another embodiment the steam treatment is used to remove most of thehalogen from the formed catalyst. Prior heating in a low moistureenvironment is not necessary. The formed catalyst is steam-treated at atemperature from about 400° C. to about 1000° C., preferably from about500° C. to about 800° C., more preferably from about 600° C. to about700° C. The period during which the catalysts is heated in unit 20ranges from 0.1 to 5 hours, preferably from 0.25 to 4 hours. The steamtreatment may remove from about 70% to about 99.99% by weight,preferably from about 95% to about 99.99% by weight, more preferablyfrom about 98% to about 99.99% by weight, of the chlorine in the formedcatalyst.

Following the steam treatment, the catalyst may be oxygen treated toremove template material that may have remained in the pores of thesieve. The calcination temperature of this oxygen environment may be thesame or different than the temperature of the steam treatment.

In another embodiment, steam treatment of the catalyst may take placeafter an oxygen heat treatment. The catalyst is heated in an oxygenenvironment at a temperature from 400° C. to 1000° C., preferably fromabout 500° C. to about 800° C., more preferably from about 600° C. toabout 700° C. The period during which the catalysts is heated in unit 20ranges from 0.1 to 5 hours, preferably from 0.25 to 4 hours.Approximately 50% to 95% by weight, preferably 75% to 95% by weight, ofthe chlorine in the formed catalyst is removed during the oxygen heattreatment. The oxygen treated catalyst is then contacted with steam toremove additional amounts of chlorine from the catalyst. This steamcontacted catalyst will contain about 10 ppmw to about 600 ppmwchlorine, preferably 10 ppmw to about 200 ppmw chlorine, more preferably10 ppmw to about 80 ppmw chlorine.

The oxygen heat treatment and the steam treatment of the catalyst maytake place in separate heating units or in the same heating unit thoughin different regions of that unit. For example, the oxygen environmentmay be introduced near the entrance to the heating unit and steam addednear the middle of the heating unit. In this way partial calcination ofthe catalyst occurs prior to the catalyst contacting the steam.

The catalysts of the present invention are useful in a variety ofprocesses including: cracking, hydrocracking, isomerization,polymerisation, reforming, hydrogenation, dehydrogenation, dewaxing,hydrodewaxing, absorption, alkylation, transalkylation, dealkylation,hydrodecylization, disproportionation, oligomerization,dehydrocyclization and combinations thereof. Due to the low level, oreven absence of halogen in the non zeolitic catalyst, less corrosiveacid is created during the catalytic process. Reactors may thus be usedfor longer periods of time before repair or replacement needs to takeplace.

The preferred processes of the present invention include a processdirected to the conversion of a feedstock comprising one or moreoxygenates to one or more olefin(s) and a process directed to theconversion of ammonia and one or more oxygenates to alkyl amines and inparticular methylamines.

In a preferred embodiment of the process of the invention, the feedstockcontains one or more oxygenates, more specifically, one or more organiccompound(s) containing at least one oxygen atom. In the most preferredembodiment of the process of the invention, the oxygenate in thefeedstock is one or more alcohol(s), preferably aliphatic alcohol(s)where the aliphatic moiety of the alcohol(s) has from 1 to 20 carbonatoms, preferably from 1 to 10 carbon atoms, and most preferably from 1to 4 carbon atoms. The alcohols useful as feedstock in the process ofthe invention include lower straight and branched chain aliphaticalcohols and their unsaturated counterparts.

Non-limiting examples of oxygenates include methanol, ethanol,n-propanol, isopropanol, methyl ethyl ether, dimethyl ether, diethylether, di-isopropyl ether, formaldehyde, dimethyl carbonate, dimethylketone, acetic acid, and mixtures thereof.

In the most preferred embodiment, the feedstock is selected from one ormore of methanol, ethanol, dimethyl ether, diethyl ether or acombination thereof, more preferably methanol and dimethyl ether, andmost preferably methanol.

In the most preferred embodiment, the feedstock, preferably of one ormore oxygenates, is converted in the presence of a catalyst intoolefin(s) having 2 to 6 carbons atoms, preferably 2 to 4 carbon atoms.Most preferably, the olefin(s), alone or combination, are converted froma feedstock containing an oxygenate, preferably an alcohol, mostpreferably methanol, to the preferred olefin(s) ethylene and/orpropylene.

The most preferred process is generally referred to as gas-to-olefins(GTO) or alternatively, methanol-to-olefins (MTO). In a MTO process,typically an oxygenated feedstock, most preferably a methanol containingfeedstock, is converted in the presence of a catalyst into one or moreolefin(s), preferably and predominantly, ethylene and/or propylene,often referred to as light olefin(s).

In one embodiment of the process for conversion of a feedstock,preferably a feedstock containing one or more oxygenates, the amount ofolefin(s) produced based on the total weight of hydrocarbon produced isgreater than 50 weight percent, preferably greater than 60 weightpercent, more preferably greater than 70 weight percent.

The feedstock, in one embodiment, contains one or more diluent(s),typically used to reduce the concentration of the feedstock, and aregenerally non-reactive to the feedstock or catalyst. Non-limitingexamples of diluents include helium, argon, nitrogen, carbon monoxide,carbon dioxide, water, essentially non-reactive paraffins (especiallyalkanes such as methane, ethane, and propane), essentially non-reactivearomatic compounds, and mixtures thereof. The most preferred diluentsare water and nitrogen, with water being particularly preferred.

The diluent, water, is used either in a liquid or a vapour form, or acombination thereof. The diluent is either added directly to a feedstockentering into a reactor or added directly into a reactor, or added witha molecular sieve catalyst composition. In one embodiment, the amount ofdiluent in the feedstock is in the range of from about 1 to about 99mole percent based on the total number of moles of the feedstock anddiluent, preferably from about 1 to 80 mole percent, more preferablyfrom about 5 to about 50, most preferably from about 5 to about 25. Inone embodiment, other hydrocarbons are added to a feedstock eitherdirectly or indirectly, and include olefin(s), paraffin(s), aromatic(s)(see for example U.S. Pat. No. 4,677,242, addition of aromatics) ormixtures thereof, preferably propylene, butylene, pentylene, and otherhydrocarbons having 4 or more carbon atoms, or mixtures thereof.

The process for converting a feedstock, especially a feedstockcontaining one or more oxygenates, in the presence of a catalyst of theinvention, is carried out in a reaction process in a reactor, where theprocess is a fixed bed process, a fluidised bed process (includes aturbulent bed process), preferably a continuous fluidised bed process,and most preferably a continuous high velocity fluidised bed process.

The reaction processes can take place in a variety of catalytic reactorssuch as hybrid reactors that have a dense bed or fixed bed reactionzones and/or fast fluidised bed reaction zones coupled together,circulating fluidised bed reactors, riser reactors, and the like.Suitable conventional reactor types are described in for example U.S.Pat. No. 4,076,796, U.S. Pat. No. 6,287,522 (dual riser), andFluidization Engineering, D. Kunii and O. Levenspiel, Robert E. KriegerPublishing Company, New York, N.Y. 1977, which are all herein fullyincorporated by reference.

The preferred reactor type are riser reactors generally described inRiser Reactor, Fluidization and Fluid-Particle Systems, pages 48 to 59,F. A. Zenz and D. F. Othmo, Reinhold Publishing Corporation, New York,1960, and U.S. Pat. No. 6,166,282 (fast-fluidised bed reactor), and U.S.patent application Ser. No. 09/564,613 filed May 4, 2000 (multiple riserreactor), which are all herein fully incorporated by reference.

In the preferred embodiment, a fluidised bed process or high velocityfluidised bed process includes a reactor system, a regeneration systemand a recovery system.

The reactor system preferably is a fluid bed reactor system having afirst reaction zone within one or more riser reactor(s) and a secondreaction zone within at least one disengaging vessel, preferablycomprising one or more cyclones. In one embodiment, the one or moreriser reactor(s) and disengaging vessel is contained within a singlereactor vessel. Fresh feedstock, preferably containing one or moreoxygenates, optionally with one or more diluent(s), is fed to the one ormore riser reactor(s) in which a catalyst or coked version thereof isintroduced. In one embodiment, the catalyst or coked version thereof iscontacted with a liquid or gas, or combination thereof, prior to beingintroduced to the riser reactor(s), preferably the liquid is water ormethanol, and the gas is an inert gas such as nitrogen.

In an embodiment, the amount of fresh feedstock fed separately orjointly with a vapour feedstock, to a reactor system is in the range offrom 0.1 weight percent to about 85 weight percent, preferably fromabout 1 weight percent to about 75 weight percent, more preferably fromabout 5 weight percent to about 65 weight percent based on the totalweight of the feedstock including any diluent contained therein. Theliquid and vapour feedstocks are preferably the same composition, orcontain varying proportions of the same or different feedstock with thesame or different diluent.

The feedstock entering the reactor system is preferably converted,partially or fully, in the first reactor zone into a gaseous effluentthat enters the disengaging vessel along with a coked catalyst. In thepreferred embodiment, cyclone(s) within the disengaging vessel aredesigned to separate the catalyst, preferably a coked catalyst, from thegaseous effluent containing one or more olefin(s) within the disengagingzone. Cyclones are preferred, however, gravity effects within thedisengaging vessel will also separate the catalyst compositions from thegaseous effluent. Other methods for separating the catalyst compositionsfrom the gaseous effluent include the use of plates, caps, elbows, andthe like.

In one embodiment of the disengaging system, the disengaging systemincludes a disengaging vessel; typically a lower portion of thedisengaging vessel is a stripping zone. In the stripping zone the cokedcatalyst is contacted with a gas, preferably one or a combination ofsteam, methane, carbon dioxide, carbon monoxide, hydrogen, or an inertgas such as argon, preferably steam, to recover adsorbed hydrocarbonsfrom the coked catalyst that is then introduced to the regenerationsystem. In another embodiment, the stripping zone is in a separatevessel from the disengaging vessel and the gas is passed at a gas hourlysuperficial velocity (GHSV) of from 1 hr⁻¹ to about 20,000 hr⁻¹ based onthe volume of gas to volume of coked catalyst, preferably at an elevatedtemperature from 250° C. to about 750° C., preferably from about 350° C.to 650° C., over the coked catalyst.

The conversion temperature employed in the conversion process,specifically within the reactor system, is in the range of from about200° C. to about 1000° C., preferably from about 250° C. to about 800°C., more preferably from about 250° C. to about 750 ° C., yet morepreferably from about 300° C. to about 650° C., yet even more preferablyfrom about 350° C. to about 600° C. most preferably from about 350° C.to about 550° C.

The conversion pressure employed in the conversion process, specificallywithin the reactor system, varies over a wide range including autogenouspressure. The conversion pressure is based on the partial pressure ofthe feedstock exclusive of any diluent therein. Typically the conversionpressure employed in the process is in the range of from about 0.1 kPaato about 5 MPaa, preferably from about 5 kPaa to about 1 MPaa, and mostpreferably from about 20 kPaa to about 500 kpaa.

The weight hourly space velocity (WHSV), particularly in a process forconverting a feedstock containing one or more oxygenates in the presenceof a catalyst within a reaction zone, is defined as the total weight ofthe feedstock excluding any diluents to the reaction zone per hour perweight of molecular sieve in the catalyst in the reaction zone. The WHSVis maintained at a level sufficient to keep the catalyst composition ina fluidised state within a reactor.

Typically, the WHSV ranges from about 1 hr⁻¹ to about 5000 hr⁻¹,preferably from about 2 hr⁻¹ to about 3000 hr⁻¹, more preferably fromabout 5 hr⁻¹ to about 1500 hr⁻¹, and most preferably from about 10 hr⁻¹to about 1000 hr⁻¹. In one preferred embodiment, the WHSV is greaterthan 20 hr⁻¹; preferably the WHSV for conversion of a feedstockcontaining methanol and dimethyl ether is in the range of from about 20hr⁻¹ to about 300 hr⁻¹.

The superficial gas velocity (SGV) of the feedstock including diluentand reaction products within the reactor system is preferably sufficientto fluidise the catalyst within a reaction zone in the reactor. The SGVin the process, particularly within the reactor system, moreparticularly within the riser reactor(s), is at least 0.1 meter persecond (m/sec), preferably greater than 0.5 m/sec, more preferablygreater than 1 m/sec, even more preferably greater than 2 m/sec, yeteven more preferably greater than 3 m/sec, and most preferably greaterthan 4 m/sec. See for example U.S. patent application Ser. No.09/708,753 filed Nov. 8, 2000, which is herein incorporated byreference.

In one preferred embodiment of the process for converting an oxygenateto olefin(s) using a silicoaluminophosphate catalyst, the process isoperated at a WHSV of at least 20 hr⁻¹ and a Temperature CorrectedNormalized Methane Selectivity (TCNMS) of less than 0.016, preferablyless than or equal to 0.01. See for example U.S. Pat. No. 5,952,538,which is herein fully incorporated by reference.

In another embodiment of the processes for converting an oxygenate suchas methanol to one or more olefin(s) using a catalyst, the WHSV is from0.01 hr⁻¹ to about 100 hr⁻¹, at a temperature of from about 350° C. to550° C., and silica to Me₂O₃ (Me is a Group IIIA or VIII element fromthe Periodic Table of Elements) molar ratio of from 300 to 2500. See forexample EP-0 642 485 B1, which is herein fully incorporated byreference.

Other processes for converting an oxygenate such as methanol to one ormore olefin(s) using a catalyst are described in PCT WO 01/23500published Apr. 5, 2001 (propane reduction at an average catalystfeedstock exposure of at least 1.0), which is herein incorporated byreference.

The coked catalyst is withdrawn from the disengaging vessel, preferablyby one or more cyclones(s), and introduced to the regeneration system.The regeneration system comprises a regenerator where the coked catalystcomposition is contacted with a regeneration medium, preferably a gascontaining oxygen, under general regeneration conditions of temperature,pressure and residence time.

Non-limiting examples of the regeneration medium include one or more ofoxygen, O₃, SO₃, N₂O, NO, NO₂, N₂O₅, air, air diluted with nitrogen orcarbon dioxide, oxygen and water (U.S. Pat. No. 6,245,703), carbonmonoxide and/or hydrogen. The regeneration conditions are those capableof burning coke from the coked catalyst composition, preferably to alevel less than 0.5 weight percent based on the total weight of thecoked catalyst entering the regeneration system. The coked catalystwithdrawn from the regenerator forms a regenerated catalyst.

The regeneration temperature is in the range of from about 200° C. toabout 1500° C., preferably from about 300° C. to about 1000° C., morepreferably from about 450° C. to about 750° C., and most preferably fromabout 550° C. to 700° C. The regeneration pressure is in the range offrom about 15 psia (103 kPaa) to about 500 psia (3448 kPaa), preferablyfrom about 20 psia (138 kPaa) to about 250 psia (1724 kpaa), morepreferably from about 25 psia (172 kPaa) to about 150 psia (1034 kPaa),and most preferably from about 30 psia (207 kPaa) to about 60 psia (414kPaa).

The preferred residence time of the catalyst in the regenerator is inthe range of from about one minute to several hours, most preferablyabout one minute to 100 minutes, and the preferred volume of oxygen inthe gas is in the range of from about 0.01 mole percent to about 5 molepercent based on the total volume of the gas.

In one embodiment, regeneration promoters, typically metal containingcompounds such as platinum, palladium and the like, are added to theregenerator directly, or indirectly, for example with the coked catalystcomposition. Also, in another embodiment, a fresh catalyst is added tothe regenerator containing a regeneration medium of oxygen and water asdescribed in U.S. Pat. No. 6,245,703, which is herein fully incorporatedby reference.

In an embodiment, a portion of the coked catalyst from the regeneratoris returned directly to the one or more riser reactor(s), or indirectly,by pre-contacting with the feedstock, or contacting with fresh catalyst,or contacting with a regenerated catalyst or a cooled regeneratedcatalyst described below.

The burning of coke is an exothermic reaction, and in an embodiment, thetemperature within the regeneration system is controlled by varioustechniques in the art including feeding a cooled gas to the regeneratorvessel, operated either in a batch, continuous, or semi-continuous mode,or a combination thereof. A preferred technique involves withdrawing theregenerated catalyst from the regeneration system and passing theregenerated catalyst through a catalyst cooler that forms a cooledregenerated catalyst. The catalyst cooler, in an embodiment, is a heatexchanger that is located either internal or external to theregeneration system.

In one embodiment, the cooler regenerated catalyst is returned to theregenerator in a continuous cycle, alternatively, (see U.S. patentapplication Ser. No. 09/587,766 filed Jun. 6, 2000) a portion of thecooled regenerated catalyst is returned to the regenerator vessel in acontinuous cycle, and another portion of the cooled molecular sieveregenerated catalyst is returned to the riser reactor(s), directly orindirectly, or a portion of the regenerated catalyst or cooledregenerated catalyst is contacted with by-products within the gaseouseffluent (PCT WO 00/49106 published Aug. 24, 2000), which are all hereinfully incorporated by reference. In another embodiment, a regeneratedcatalyst contacted with an alcohol, preferably ethanol, 1-propnaol,1-butanol or mixture thereof, is introduced to the reactor system, asdescribed in U.S. patent application Ser. No. 09/785,122 filed Feb. 16,2001, which is herein fully incorporated by reference.

Other methods for operating a regeneration system are in disclosed U.S.Pat. No. 6,290,916 (controlling moisture), which is herein fullyincorporated by reference.

The regenerated catalyst withdrawn from the regeneration system,preferably from the catalyst cooler, is combined with a fresh catalystand/or re-circulated catalyst and/or feedstock and/or fresh gas orliquids, and returned to the riser reactor(s). In another embodiment,the regenerated catalyst withdrawn from the regeneration system isreturned to the riser reactor(s) directly, preferably after passingthrough a catalyst cooler. In one embodiment, a carrier, such as aninert gas, feedstock vapour, steam or the like, semi-continuously orcontinuously, facilitates the introduction of the regenerated catalystto the reactor system, preferably to the one or more riser reactor(s).

By controlling the flow of the regenerated catalyst or cooledregenerated catalyst from the regeneration system to the reactor system,the optimum level of coke on the catalyst entering the reactor ismaintained. There are many techniques for controlling the flow of acatalyst described in Michael Louge, Experimental Techniques,Circulating Fluidised Beds, Grace, Avidan and Knowlton, eds. Blackie,1997 (336–337), which is herein incorporated by reference.

Coke levels on the catalyst are measured by withdrawing from theconversion process the catalyst at a point in the process anddetermining its carbon content. Typical levels of coke on the catalyst,after regeneration is in the range of from 0.01 weight percent to about15 weight percent, preferably from about 0.1 weight percent to about 10weight percent, more preferably from about 0.2 weight percent to about 5weight percent, and most preferably from about 0.3 weight percent toabout 2 weight percent based on the total weight of the molecular sieveand not the total weight of the catalyst.

In one preferred embodiment, the mixture of fresh catalyst andregenerated catalyst and/or cooled regenerated catalyst contains in therange of from about 1 to 50 weight percent, preferably from about 2 to30 weight percent, more preferably from about 2 to about 20 weightpercent, and most preferably from about 2 to about 10 coke orcarbonaceous deposit based on the total weight of the mixture ofcatalysts. See for example U.S. Pat. No. 6,023,005, which is hereinfully incorporated by reference.

The gaseous effluent is withdrawn from the disengaging system and ispassed through a recovery system. There are many well-known recoverysystems, techniques and sequences that are useful in separatingolefin(s) and purifying olefin(s) from the gaseous effluent. Recoverysystems generally comprise one or more or a combination of a variousseparation, fractionation and/or distillation towers, columns,splitters, or trains, reaction systems such as ethylbenzene manufacture(U.S. Pat. No. 5,476,978) and other derivative processes such asaldehydes, ketones and ester manufacture (U.S. Pat. No. 5,675,041), andother associated equipment for example various condensers, heatexchangers, refrigeration systems or chill trains, compressors,knock-out drums or pots, pumps, and the like.

The metalloaluminophosphate molecular sieve materials and catalystcompositions of the present invention may be used in the manufacture ofalkylamines, using ammonia. Examples of suitable processes are asdescribed in published European Patent Application EP 0 993 867 A1, andin U.S. Pat. No. 6,153,798 to Hidaka et. al, which are herein fullyincorporated by reference.

This invention will be better understood with reference to the followingexamples, which are intended to illustrate specific embodiments withinthe overall scope of the invention as claimed.

EXAMPLE 1

SAPO-34 molecular sieve, 50% by weight, aluminum chlorhydrol, 10% byweight, and UF grade kaolin clay, 40% by weight, was mixed withsufficient water to produce a slurry with approximately 40% by weightsolids. The slurry was fed into a spray drier to form spray driedcatalyst. The spray dried catalyst was analyzed by XRF (X-rayFluorescence) spectroscopy. The amount of chlorine in the spray driedcatalyst was 33,800 ppmw. The GAL Index of the un-calcined catalyst wasgreater than 50.

EXAMPLES 2–4

Spray dried catalyst of Example 1 was heated in a nitrogen stream attemperatures of 600° C., 650° C. and 700° C. for one hour. The heattreated catalyst was then analyzed by XRF to determine the amount ofresidual chlorine remaining in the catalyst. Table 1 lists the residualchlorine content of each catalyst.

EXAMPLES 5–7

Spray dried catalyst of Example 1 was heated in a nitrogen stream attemperatures of 600° C., 650° C. and 700° C. for nine hours. The heattreated catalysts were then analyzed by XRF to determine the amount ofresidual chlorine remaining in each catalyst. Table 1 lists the residualchlorine content of each catalyst.

EXAMPLES 8

Spray dried catalyst of Example 1 was heated in a nitrogen stream attemperatures of 650° C. for five hours followed by heating in air at650° C. for two hours. The heat treated catalyst was then analyzed byXRF to determine the amount of residual chlorine remaining in thecatalyst. The chlorine content of the catalyst was 390 ppm by weight.

EXAMPLES 9–11

The spray dried catalyst of Example 1 was heated in a nitrogen stream attemperatures of 600° C., 650° C. and 700° C. for one hour followed byheating in air at 600° C., 650° C. and 700° C. for one hour,respectively. The heat treated catalysts were then analyzed by XRF todetermine the amount of residual chlorine remaining in each catalyst.Table 1 lists the residual chlorine content of each catalyst.

TABLE 1 Example Temperature, Sweep gas Sweep gas Chlorine, No. ° C.time, hrs time, hrs ppmw 1 n/a n/a n/a 33,800 2 600 N2/1 N/A 620 3 650N2/1 N/A 520 4 700 N2/1 N/A 480 5 600 N2/9 N/A 440 6 650 N2/9 N/A 430 7700 N2/9 N/A 350 8 650 N2/5 air/2 390 9 600 N2/1 air/1 510 10  650 N2/1air/1 470 11  700 N2/1 air/1 430

EXAMPLE 12

The spray dried catalyst was heated at 600° C. in air for 120 minutes inan open container placed in an electrically heated muffle furnace. Thecalcined catalyst contained 1090 ppm chlorine (see Table 2).

EXAMPLE 13

The spray dried catalyst was heated at 650° C. in air for 120 minutes inan open container placed in an electrically heated muffle furnace. Thecalcined catalyst contained 730 ppm chlorine, and the GAL Index was 1.85(see Table 2).

EXAMPLE 14

The spray dried catalyst was heated at 700° C. in air for 120 minutes inan open container placed in an electrically heated muffle furnace. Thecalcined catalyst contained 350 ppm chlorine (see Table 2).

EXAMPLE 15

The spray dried catalyst was heated at 600° C. in air for 120 minutes inan open container placed in an electrically heated muffle furnace. Thecalcined catalyst, 12 g, was placed in a ¾″ OD stainless steel, packedbed tubular reactor that was electrically heated. About 1 g/min of steamwas fed to the reactor maintained at a temperature of 600° C. Thecatalyst was heated in the presence of steam for 120 minutes. Thechlorine content of the treated catalyst was 250 ppm (see Table 2).

EXAMPLE 16

The same procedure as in Example 15 was used except that the temperaturewas maintained at 650° C. for both the heating in air and heating insteam. The chlorine content of the treated catalyst was 140 ppm, and theGAL Index was 1.48 (see Table 2).

EXAMPLE 17

The same procedure as in Example 15 was used except that the temperaturewas maintained at 700° C. for both the heating in air and heating insteam. The chlorine content of the treated catalyst was 30 ppm (seeTable 2).

EXAMPLE 18

The same procedure as in Examples 12 was used except that the catalystwas heated in air for 240 minutes. The chlorine content of the treatedcatalyst was 830 ppm (see Table 2).

EXAMPLE 19

The same procedure as in Examples 13 was used except that the catalystwas heated in air for 240 minutes. The chlorine content of the treatedcatalyst was 590 ppm (see Table 2).

EXAMPLE 20

The same procedure as in Examples 14 was used except that the catalystwas heated in air for 240 minutes. The chlorine content of the treatedcatalyst was 290 ppm (see Table 2).

EXAMPLES 21–26

The same procedure as in Examples 15 were used except the times andtemperatures of heating in air and the times and temperatures of heatingin steam as indicated in Table 2.

EXAMPLE 27

The spray dried catalyst was heated at 600° C. in air for 120 minutes inan open container placed in an electrically heated muffle furnace. Thecalcined catalyst then placed in a ¾″ OD stainless steel, packed bedtubular reactor that was electrically heated. About 1 g/min of steam atabout 1 atm was fed to the reactor maintained at a temperature of 600°C. The catalyst was heated in the presence of steam for 240 minutes. Thechlorine content of the treated catalyst was 150 ppmw, and the GAL Indexwas 2.24 (see Table 2).

EXAMPLE 28

The same procedure as in Example 27 was used except the temperatures ofheating in air and the steam treatment was 650° C. The chlorine contentof the treated catalyst was 40 ppmw, and the GAL Index was 1.62 (seeTable 2).

As summarized in Table 2, heating in air for 120 minutes at 600° C.,650° C. and 700° C. without a subsequent steam treatment reduces thechlorine content to 1090, 730 or 350 ppm respectively. Increasing theheating time to 240 minutes at 600° C., 650° C. and 700° C., results inthe further reduction in chlorine content to 830, 590 or 290 ppm,respectively. As indicated only small amounts of additional chlorine isremoved by a doubling of the heating time. For example, heating at 650°C. during the first 120 minutes reduces the chlorine content in thecatalyst by about 98%, i.e., from 33,800 ppm to 730 ppm. Heating for asecond 120 minutes reduces the remaining chlorine content by anadditional 19%, i.e., from 730 ppm to 590 ppm.

Heating in air for 120 minutes followed by heating in the presence ofsteam for 120 minutes at temperatures of 600° C., 650° C., and 700° C.reduces the chlorine content to 250, 140 and 30 ppm, respectively.Increasing the time the catalyst is heated in air and steam to 240minutes, respectively, has little affect on further reducing thechlorine content as shown by a comparison of Examples 15–17 withExamples 21–23, respectively.

Examples 24–26 indicate that increasing the time the catalyst is steamtreated at a given temperature (650° C. in these examples) following theheat treatment in air for 120 minutes results in a yet greater reductionin chlorine content. The most dramatic reduction in chlorine content ismade during the first 15 minutes of contacting the heat treated catalystwith steam. For example, comparison of Example 13 with Example 24suggests that the chlorine content is reduced from 730 ppm to 230 ppmafter an additional 15 minute steam treatment at 650° C. This amounts toan additional chlorine reduction of about 68%. Also, as indicated inTable 2 greater than 99% of the chlorine may be removed from thecatalyst following the steam treatment of the catalyst.

TABLE 2 Time Time Example (air) (steam) GAL Chlorine No. Temp. ° C. min.min. Index Ppmw 1 N/A N/A N/A >50 33800 12 600 120 0 1090 13 650 120 01.85 730 14 700 120 0 350 15 600 120 120 250 16 650 120 120 1.48 140 17700 120 120 30 18 600 240 0 830 19 650 240 0 590 20 700 240 0 290 21 600240 240 220 22 650 240 240 160 23 700 240 240 40 24 650 120 15 230 25650 120 30 120 26 650 120 60 70 27 600 120 240 2.24 150 28 650 120 2401.62 40

The attrition properties of Examples 1, 13, 16, 27, and 28 are listed inTable 2. Attrition properties of catalysts can be defined by the GrossAttrition Loss (GAL) Index. The smaller the GAL Index the more resistantto attrition is the catalyst. The GAL Index is measured in the followingmanner. About 6.0±0.1 g of SAPO catalyst was added to an attrition cupof an attrition apparatus known in the art. 23,700 scc/min of nitrogengas was bubbled through a water-containing bubbler to humidify the N₂.The wet nitrogen passed through the attrition cup and exited theattrition apparatus through a porous fiber thimble. This thimbleseparates the fine catalyst particles resulting from the attrition ofthe catalyst particles in the attrition cup as the catalyst particlesare circulated in the attrition cup by the fast flowing nitrogen gas.The pore size of the thimble determines the size of the fine particlesthat are separated from the catalyst. The pore size of the thimble usedto measure the GAL Index was less than about 2 μm.

The nitrogen flow passing through the attrition cup was maintained for60 minutes. The contents of the attrition cup were transferred to anelutriation cup. The elutriation cup is designed not to cause furtherattrition of the catalyst particles, but to remove any fine particlesremaining in the attrition cup so that the fine particles may beincluded in the GAL Index. 23,700 scc/min of nitrogen gas was passedthrough the elutriation cup for 30 minutes. Additional fine particleswere separated by the thimble. The collection of fine SAPO particlesseparated by the thimble were weighed. The amount in grams of fineparticles divided by the original amount of catalyst added to theattrition cup is the GAL Index.GAL Index=C/(B+C)×100wherein

-   -   B=weight of catalyst in elutriation cup    -   C=weight of collected fine catalyst particles

Having now fully described this invention, it will be appreciated bythose skilled in the art that the invention can be performed within awide range of parameters within what is claimed, without departing fromthe spirit and scope of the invention.

1. A conversion process selected from the group consisting of oxygenateconversion to olefins, oxygenate and ammonia conversion to alkylamines,and hydrocarbon conversion comprising cracking, hydrocracking,isomerization, polymerisation, reforming, hydrogenation,dehydrogenation, dewaxing, hydrodewaxing, alkylation, transalkylation,dealkylation, hydrodecyclization, disproportionation, oligomerization,dehydrocyclization or combinations thereof, the process comprising thesteps of: (a) introducing a feedstock to a reactor system in thepresence of a catalyst comprising a non zeolitic molecular sieve,inorganic oxide matrix, and matrix material, wherein the catalystcontains from 10 ppm to 600 ppm by weight halogen; to thereby carry outa reaction process which converts said feedstock into a gaseous effluentstream containing one or more conversion products; (b) withdrawing fromthe reactor system said gaseous effluent stream; and (c) passing theeffluent gas through a recovery system recovering at least the one ormore conversion products.
 2. The process of claim 1 wherein thefeedstock comprises one or more oxygenates.
 3. The process of claim 2wherein the one or more oxygenates comprises methanol.
 4. The process ofclaim 1 wherein the one or more conversion products comprises one ormore olefins.
 5. The process of claim 4 wherein the one or more olefinscomprises ethylene, propylene and mixtures thereof.
 6. The process ofclaim 1 wherein the feedstock comprises one or more oxygenates andammonia.
 7. The process of claim 6 wherein the one or more conversionproducts comprises one or more alkylamines.
 8. The process of claim 7wherein the one or more alkylamines comprises one or more methylamines.9. The process of claim 6 wherein the one or more oxygenates comprisesmethanol.
 10. The process of claim 1 wherein the feedstock comprises oneor more oxygenates and one or more aromatic compound.
 11. The process ofclaim 10 wherein the one or more conversion products comprises one ormore alkylated aromatic compound.
 12. The process of claim 1 wherein theconversion process is cracking.
 13. The process of claim 1 wherein theconversion process is dewaxing.
 14. The process of claim 1 wherein thehalogen is chlorine.
 15. The process of claim 14 wherein die catalystcontains from 10 ppmw to 400 ppmw chlorine.
 16. The process of claim 14wherein the catalyst contains from 10 ppmw to 200 ppmw chlorine.
 17. Theprocess of claim 15 wherein the catalyst contains from 10 ppmw to 80ppmw chlorine.
 18. The process of claim 1 wherein the catalyst has a GALIndex of less than
 5. 19. The process of claim 18 wherein the GAL Indexis less than
 3. 20. The process of claim 1 wherein the non-zeoliticmolecular sieve is selected from SAPO-5, SAPO-8, SAPO-11, SAPO-16,SAPO-17, SAPO-18, SAPO-20, SAPO-31, SAPO)-34, SAPO-35, SAPO-36, SAPO-37,SAPO-40, SAPO-41, SAPO-42, SAPO-44, SAPO-47, SAPO-56, ALPO-5, ALPO-11,ALPO-18, ALPO-31, ALPO-34, ALPO-36, ALPO-37, ALPO-46, the metalcontaining forms of each thereof, and mixtures thereof.
 21. The processof claim 1 wherein the catalyst comprises 20% to 45% by weightnon-zeolitic molecular sieve, 5% to 20% by weight of inorganic oxidematrix, and 20% to 70% by weight matrix material.
 22. The process ofclaim 21 wherein the catalyst comprises 25% to 42% by weightnon-zeolitic molecular sieve, 8% to 15% by weight of inorganic oxidematrix, and 40% to 60% by weight matrix material.
 23. The process ofclaim 19 wherein the GAL Index is less than
 2. 24. The process of claim1 wherein the inorganic oxide matrix comprises an aluminum oxide matrix.25. The process of claim 2 wherein the source of chlorine is aluminumchlorhydrol.
 26. A conversion process selected from the group consistingof oxygenate conversion to olefins, oxygenate and ammonia conversion toalkylamines, and hydrocarbon conversion comprising cracking,hydrocracking, isomerization, polymerisation, reforming, hydrogenation,dehydrogenation, dewaxing, hydrodewaxing, alkylation, transalkylation,dealkylation, hydrodecyclization, disproportionation, oligomerization,dehydrocyclization or combinations thereof, the process comprising thesteps of: (a) providing a halogen-containing non-zeolitic molecularsieve catalyst; (b) heating the catalyst in a low moisture environmentat a temperature from 400° C. to 1000° C.; (c) contacting the heatedcatalyst with steam at a temperature from 400° C. to 1000° C. to producea steam-treated catalyst; (d) introducing a feedstock to a reactorsystem in the presence of the steam-treated catalyst obtained at step(c); to thereby carry out a reaction process which converts saidfeedstock into a gaseous effluent stream containing one or moreconversion products; (e) withdrawing from the reactor system saidgaseous effluent stream; and (f) passing the effluent gas through arecovery system recovering at least the one or more conversion products.27. The process of claim 26 wherein the halogen is chlorine.
 28. Theprocess of claim 26 wherein the feedstock comprises one or moreoxygenates.
 29. The process of claim 28 wherein the one or moreoxygenates comprises methanol.
 30. The process of claim 26 wherein theone or more conversion products comprises one or more olefins.
 31. Theprocess of claim 30 wherein the one or more olefins comprises ethylene,propylene and mixtures thereof.
 32. The process of claim 26 wherein thefeedstock comprises one or more oxygenates and ammonia.
 33. The processof claim 32 wherein the one or more conversion products comprises one ormore alkylamines.
 34. The process of claim 33 wherein the one or morealkylamines comprises one or more methylamines.
 35. The process of claim32 wherein the one or more oxygenates comprises methanol.
 36. Theprocess of claim 26 wherein the feedstock comprises one or moreoxygenates and one or more aromatic compound.
 37. The process of claim36 wherein the one or more conversion products comprises one or morealkylated aromatic compound.
 38. The process of claim 26 wherein theconversion process is cracking.
 39. The process of claim 26 wherein theconversion process is dewaxing.
 40. A conversion process selected fromthe group consisting of oxygenate conversion to olefins, oxygenate andammonia conversion to alkylamines, and hydrocarbon conversion comprisingcracking, hydrocracking, isomerization, polymerisation, reforming,hydrogenation, dehydrogenation, dewaxing, hydrodewaxing, alkylation,transalkylation, dealkylation, hydrodecyclization, disproportionation,oligomerization, dehydrocyclization or combinations thereof, the processcomprising the steps of: (a) providing a halogen-containing non-zeoliticmolecular sieve catalyst; (b) heating the catalyst in an oxygenenvironment at a temperature from 400° C. to 1000° C.; (c) contactingthe heated catalyst with steam at a temperature from 400° C. to 1000° C.to produce a steam-treated catalyst; (d) introducing a feedstock to areactor system in the presence of the steam-treated catalyst obtained atstep (c); to thereby carry out a reaction process which converts saidfeedstock into a gaseous effluent stream containing one or moreconversion products; (e) withdrawing from the reactor system saidgaseous effluent stream; and (f) passing the effluent gas through arecovery system recovering at least the one or more conversion products.41. The process of claim 40 wherein the halogen is chlorine.
 42. Theprocess of claim 40 wherein the feedstock comprises one or moreoxygenates.
 43. The process of claim 42 wherein the one or moreoxygenates comprises methanol.
 44. The process of claim 40 wherein theone or more conversion products comprises one or more olefins.
 45. Theprocess of claim 44 wherein the one or more olefins comprises ethylene,propylene and mixtures thereof.
 46. The process of claim 40 wherein thefeedstock comprises one or more oxygenates and ammonia.
 47. The processof claim 46 wherein the one or more conversion products comprises one ormore alkylamines.
 48. The process of claim 47 wherein the one or morealkylamines comprises one or more methylamines.
 49. The process of claim46 wherein the one or more oxygenates comprises methanol.
 50. Theprocess of claim 40 wherein the feedstock comprises one or moreoxygenates and one or more aromatic compound.
 51. The process of claim50 wherein the one or more conversion products comprises one or morealkylated aromatic compound.
 52. The process of claim 40 wherein theconversion process is cracking.
 53. The process of claim 40 wherein theconversion process is dewaxing.
 54. A conversion process selected fromthe group consisting of oxygenate conversion to olefins, oxygenate andammonia conversion to alkylamines, and hydrocarbon conversion comprisingcracking, hydrocracking, isomerization, polymerisation, reforming,hydrogenation, dehydrogenation, dewaxing, hydrodewaxing, alkylation,transalkylation, dealkylation, hydrodecyclization, disproportionation,oligomerization, dehydrocyclization or combinations thereof, the processcomprising the steps of: (a) providing a halogen-containing non-zeoliticmolecular sieve catalyst; (b) heating the catalyst in an environmentcontaining steam at a temperature from 400° C. to 1000° C.; (c) removingfrom 70% to 99.99% by weight of the halogen from the catalyst, therebyproducing a steam-treated catalyst; (d) introducing a feedstock to areactor system in the presence of the steam-treated catalyst obtained atstep (c); to thereby carry out a reaction process which converts saidfeedstock into a gaseous effluent stream containing one or moreconversion products; (e) withdrawing from the reactor system saidgaseous effluent stream; and (f) passing the effluent gas through arecovery system recovering at least the one or more conversion products.55. The process of claim 54 wherein the halogen is chlorine.
 56. Theprocess of claim 54 wherein the feedstock comprises one or moreoxygenates.
 57. The process of claim 56 wherein the one or moreoxygenates comprises methanol.
 58. The process of claim 54 wherein theone or more conversion products comprises one or more olefins.
 59. Theprocess of claim 58 wherein the one or more olefins comprises ethylene,propylene and mixtures thereof.
 60. The process of claim 54 wherein thefeedstock comprises one or more oxygenates and ammonia.
 61. The processof claim 60 wherein the one or more conversion products comprises one ormore alkylamines.
 62. The process of claim 61 wherein the one or morealkylamines comprises one or more methylamines.
 63. The process of claim60 wherein the one or more oxygenates comprises methanol.
 64. Theprocess of claim 54 wherein the feedstock comprises one or moreoxygenates and one or more aromatic compound.
 65. The process of claim64 wherein the one or more conversion products comprises one or morealkylated aromatic compound.
 66. The process of claim 54 wherein theconversion process is cracking.
 67. The process of claim 54 wherein theconversion process is dewaxing.